Integration of reforming/water splitting and electrochemical systems for power generation with integrated carbon capture

ABSTRACT

High efficiency electricity generation processes and systems with substantially zero CO2 emissions are provided. A closed looping between the unit that generates gaseous fuel (H2, CO, etc) and the fuel cell anode side is formed. In certain embodiments, the heat and exhaust oxygen containing gas from the fuel cell cathode side are also utilized for the gaseous fuel generation. The systems for converting fuel may comprise reactors configured to conduct oxidation-reduction reactions. The resulting power generation efficiencies are improved due to the minimized steam consumption for the gaseous fuel production in the fuel cell anode loop as well as the strategic mass and energy integration schemes.

This application is a divisional of U.S. patent application Ser. No.13/394,572, filed on Mar. 7, 2012 as a national stage entry under 35U.S.C. §371 of International Patent Application No. PCT/US2010/048125,filed on Sep. 8, 2010, which claims priority to U.S. ProvisionalApplication No. 61/240,508, filed on Sep. 8, 2009.

The present invention is generally directed to systems and methods ofelectricity generation with in-situ CO₂ capture. In certain embodiments,a reduction-oxidation (redox) system using one or more chemicalintermediates is utilized to convert carbonaceous fuel with CO₂ capture.This is followed by strategic integration with an electrochemicalconversion device to produce electricity. In other embodiments, watersplitting systems are integrated with the electrochemical systems.Through the process integrations, the process auxiliary powerconsumption and/or water utilization and energy used for steamgeneration are minimized.

Fossil fuels including crude oil, natural gas, and coal represent themajority of today's energy supply worldwide. The use of fossil fuels,however, requires that they be transformed to a carrier such as heat,electricity, liquid fuels, or chemicals through chemical conversionprocesses. With an increasing energy demand and concomitant concernsover the carbon emissions from fossil fuel usage, extensive efforts havebeen geared toward developing carbon neutral, efficient and economicalenergy systems that are sustainable. A transition from the use of fossilfuels to that of nuclear and renewable resources such as solar andbiomass, thus, represents the natural progression of such efforts.

Existing electricity generation technologies have one or more of thefollowing limitations/drawbacks: 1) high costs (e.g., photovoltaic,gasification, ultra-supercritical pulverized coal combustion); 2) lowefficiency (e.g., sub-critical pulverized coal combustion); 3)environmental concerns (e.g., fossil fuel power plants); and 4) safetyconcerns (e.g., nuclear power).

One of the common issues with respect to a conventional thermal powerplant is the large amount of exergy loss during cooling and reheating ofsteam. A system and method that minimizes the requirements for steamgeneration is thus desirable.

Chemical reactions between carbonaceous fuels and air/steam/CO₂ throughthe assistance of a reaction medium may represent an effective way tominimize exergy loss in the fuel conversion process. A number oftechniques have been proposed to convert carbonaceous fuels using metaloxide. For example, Watkins, U.S. Pat. No. 3,027,238, describes a methodfor producing hydrogen gas including reducing a metal oxide in areducing zone, and oxidizing the reduced metal with steam to producehydrogen in an oxidizing zone. This technique, however, is limited togaseous fuel conversion. Moreover, the gaseous fuel is only partiallyconverted by the metal oxide. Thomas, U.S. Pat. No. 7,767,191; Fan, PCTApplication No. WO 2007082089; and Fan, PCT Application No. WO2010037011 describe methods for producing hydrogen gas by reducing ametal oxide in a reduction reaction between a carbon-based fuel and ametal oxide to provide a reduced metal or metal oxide having a loweroxidation state, and oxidizing the reduced metal or metal oxide toproduce hydrogen and a metal oxide having a higher oxidation state.

Hydrogen can also be produced from water splitting throughphotoelectrolysis, thermolysis, and thermochemical routes.

To produce electricity, the aforementioned processes teach the furtherconversion of the hydrogen product in a gas turbine, gas engine, and/orfuel cell. However, a large amount of steam is used in these processesfor hydrogen generation. Simple conversion of hydrogen in conventionalhydrogen fueled power generation devices will lead to cooling andreheating of large amounts of steam/water, resulting in a largeirreversibility of the power generation system.

With increasing demand for electricity, the need arises for improvedprocesses, systems, and system components therein, which produceelectricity with higher efficiency and fewer pollutants.

Embodiments of the present invention are generally directed to highefficiency electricity generation processes and systems withsubstantially zero CO₂ emissions. A closed loop between the unit thatgenerates gaseous fuel (H₂, CO, etc.) and the fuel cell anode side isformed. In certain embodiments, the heat and exhaust oxygen containinggas from the fuel cell cathode side are also utilized for the gaseousfuel generation. The power generation efficiencies of the systemsdisclosed herein are significantly greater than state-of-the-artapproaches due to the minimized steam consumption for the gaseous fuelproduction, in the fuel cell anode loop, as well as the strategic massand energy integration schemes.

Additional features and advantages provided by embodiments of thepresent invention will be more fully understood in view of the followingdrawings and detailed description.

The following detailed description of the illustrative embodiments ofthe present invention can be best understood when read in conjunctionwith the following drawings, where like structure is indicated with likereference numerals and in which:

FIG. 1 is a general schematic illustration of a process for producingelectricity with minimal steam consumption in accordance with oneembodiment in which a fuel cell is integrated with a reforming or watersplitting system to minimize steam generation requirements in theprocess.

FIG. 2 is a general schematic illustration of an embodiment of theintegration of a fuel cell and redox based reforming/water splittingblock for electricity generation with minimal steam requirements.

FIG. 3 further illustrates the integration of an embodiment of an ironoxide redox based reforming/water splitting block and a fuel cellsystem.

FIG. 4 is a schematic of an embodiment of an iron oxide based redoxprocess using syngas derived from solid fuels such as coal or biomass asfeedstock.

FIG. 5 is a schematic of an embodiment of an iron oxide based redoxprocess using solid fuels such as coal, biomass, and/or solid wastesdirectly as feedstock and its integration with a fuel cell.

FIG. 6 is a schematic of an embodiment of a calcium basedreforming/water splitting block integrated with a fuel cell.

FIG. 7 is a schematic of an embodiment of a membrane enhancedreforming/water splitting block integrated with a fuel cell for powergeneration.

FIG. 8 is a schematic of an embodiment of a zinc based water splittingblock using solar or nuclear thermal energy and its integration with afuel cell using solar or nuclear energy.

FIG. 9 is a more detailed schematic illustrating an embodiment of anintegrated coal to electricity system using redox reactions and a solidoxide fuel cell (SOFC).

FIG. 10 illustrates an embodiment in which the unconverted fuel from thereducer of a redox based reforming/water splitting block is converted byan auxiliary fuel cell followed by an oxygen polishing step. The workingfluid between the oxidizer and the fuel cell block remains a closedloop.

Referring generally to FIG. 1, embodiments of the present invention aredirected to systems and methods for converting thermal and chemicalenergy sources into electricity with minimal steam consumption and/orauxiliary power generation and low to zero carbon emissions. Allpercentages are reported as weight percent unless otherwise noted or thecontext requires otherwise.

In one embodiment, the system is divided into two blocks or sub-systems,i.e. a reforming/water splitting block and a fuel cell block. Thereforming/water splitting block generates gaseous fuels such ashydrogen, syngas, and/or light hydrocarbons from steam/CO₂ and an energysource such as solar, nuclear, and carbonaceous fuel. The fuel cellblock converts the gaseous fuel from the reforming/water splitting blockinto electricity while generating an effluent stream which containsunconverted fuel and steam and/or CO₂, for the reforming/water splittingblock.

The steam/CO₂ effluent of the fuel cell block, which may containunconverted fuel, is recycled to the reforming/water splitting block togenerate gaseous fuel. In certain cases, minor reheating andre-pressurization of the effluent is required. Steam condensation andreheating is minimal in all cases.

To maintain the operating pressure of both the reforming/water splittingblock and the fuel cell block, a bleed of effluent and/or gaseous fuelis split from the main gaseous stream and re-pressurized. Meanwhile, are-pressurized makeup stream is merged with the main gaseous stream.Because CO₂/steam circulates between the reforming/water splitting blockand the fuel cell block along with the CO/H₂ fuel, the steam/CO₂ acts asa working fluid for electricity generation. The use of turbines, bothsteam turbines and gas turbines, is minimized in this scheme since thepartially converted gaseous fuel from the fuel cell is almost fullyrecycled to the fuel production stage. A closed loop of working fluid isformed between the reforming/water splitting block and the fuel cellblock. By minimizing the steam condensation and reheating and maximizingthe fuel conversion in the fuel cell, the irreversibility of the processis minimized.

In the case when a high temperature fuel cell such as a solid oxide fuelcell (SOFC) is used, the sulfur tolerance level is relatively high.Therefore, a simple hot gas clean up unit such as a calcium oxidesorbent bed can be integrated with the working fluid loop.

The operating pressure of the reforming/water splitting block iscomparable to the fuel cell block. Both blocks operate at pressuresbetween 1.01×10⁵ Pa and 8.11×10⁶ Pa (1 atm and 80 atm). The temperatureof the units ranges between 300° C.-1300° C. The high temperature, highpressure, spent stream from the system can be used to preheat the feedstreams, generate power, and re-pressurize the makeup stream.

The energy source for the reforming/water splitting block can either becarbonaceous fuels or thermal energy from other sources such as solar ornuclear. The carbonaceous fuels can include coal, methane, biomass,syngas, pet coke, extra heavy oil, wax and oil shale.

In the case when carbonaceous fuel is used, an oxygen carrier or CO₂sorbent is used to reform/gasify the fuel into hydrogen and/or CO In thecase when thermal energy from solar or nuclear is used, athermo-chemical water splitting scheme is used to convert thermal energyto hydrogen and oxygen.

FIG. 2 illustrates a general process configuration in which acarbonaceous fuel is indirectly reformed or gasified with steam/CO₂using a metal oxide based oxygen carrying particle. The reaction in thereduction stage is

MeO_(x)+fuel=MeO_(y)+CO₂+H₂O

In most cases, the metal oxide, the reactor design, and the operatingmode are selected such that at least 80% of the fuel is converted intoCO₂ and steam. In some cases, an oxygen polishing step is used to fullycombust the unconverted fuel (<20%) into sequestrable CO₂ and H₂O. Inpreferred embodiments, at least 95% of the fuel is converted into CO₂and steam. The exhaust gas stream from the reducing step is thussequestrable.

The reaction in the oxidation stage of FIG. 2 is

MeO_(y)+(x−y) H₂O/CO₂=MeO_(x)+(x−y) H₂/CO

The feed for the oxidation stage, directly withdrawn from the exhaust ofthe fuel cell anode side (minor recompression is conducted in certaincases), contains fuels such as H₂/CO. The fuel concentration in the fuelcell exhaust/oxidation feed typically ranges from 0 to 60%. The H₂O/CO₂in the feed stream is at least partially converted to H₂/CO, hence thefuel concentration in the gaseous stream is increased. The H₂/COconcentration in the product stream of the oxidation stage typicallyranges from 30% to 99% and is at least 5% higher than that in theexhaust stream of the fuel cell anode. The fuel enriched stream from theoxidation stage is then directly introduced back to the fuel cell forpower generation.

FIG. 3 illustrates a specific process configuration in which acarbonaceous fuel is used as the fuel and iron oxide is used as theoxygen carrier. In this embodiment, a three reactor redox system is usedto convert the fuel in a manner similar to that disclosed in Thomas,U.S. Pat. No. 7,767,191; Fan, PCT Application No. WO 2007082089; andFan, PCT Application No. WO 2010037011. The first reactor, the reducer,is configured to oxidize the carbonaceous fuel into CO₂ and steam whilereducing a metal oxide based oxygen carrier. In certain embodiments, thecarbonaceous fuel is in the form of solid particles which are suspendedby the gases in the first reaction zone until they are at least 50%converted before being elutriated towards the top of the first reactionzone. The heat required or generated in the reducer is provided orremoved by the oxygen carrier particle. The second reactor, theoxidizer, is configured to (partially) oxidize a portion of the reducedoxygen carrier with either steam or CO₂. The third reactor, thecombustor, combusts the partially oxidized oxygen carrier in theoxidizer and the remaining portion of the reduced oxygen carrier fromthe reducer using air.

The feed for the oxidizer is the exhaust from the fuel cell anode sideand the product of the oxidizer is directly used as the feed for thefuel cell anode. The oxidizer enriches the fuel content in the workingfluid (CO/H₂/CO₂/H₂O) stream. In preferred embodiments, the gaseousstream of the fuel cell anode side and the oxidizer forms a closed loopin which the addition and purging of the gaseous stream is minimal. Forexample, in certain embodiments less than 10% of the fuel rich orsteam/CO₂ rich gas stream is purged. To maintain the pressure of theworking fluid, repressurization of the fluid is performed within themain loop or a split loop. In certain embodiments, a high temperaturesorbent bed such as that using calcium oxide based sorbent is integratedinto the loop to prevent the accumulation of pollutants such as H₂S. Inother cases, sulfur treatment is carried out only on the bleed stream,the main working fluid stream is not treated.

The oxygen carrier comprises a plurality of ceramic composite particleshaving at least one metal oxide disposed on a support. Ceramic compositeparticles are described in Thomas, U.S. Pat. No. 7,767,191; Fan, PCTApplication No. WO 2007082089; and Fan, PCT Application No. WO2010037011.

Referring back to the reduction reaction in the first reactor of FIG. 3,i.e. the reducer, the reducer utilizes various carbonaceous fuels suchas syngas, methane and light hydrocarbons, coal, tars, oil shales, oilsands, tar sand, biomass, wax and coke to reduce the iron oxidecontaining ceramic composite to produce a mixture of reduced metaland/or metal oxide. The possible reduction reactions include:

FeO_(x)+Fuel→FeO_(y)+CO₂+H₂O

Fuel+CO₂→CO+H₂

Fuel+H₂O→CO+H₂

FeO_(x)+CO/H₂→FeO_(y)+CO₂/H₂O

Specifically, metallic iron (Fe) is formed in the reducer.Simultaneously, an exhaust stream that contains at least 60% CO₂ (drybasis) is produced from the reducer. In preferred schemes, the CO₂concentration exceeds 95% and is directly sequestrable.

The preferred designs for the reducer include a moving bed reactor withone or more stages, a multistage fluidized bed reactor, a step reactor,a rotatory kiln or any suitable reactors or vessels known to one ofordinary skill in the art that provide a countercurrent gas-solidcontacting pattern. The counter-current flow pattern between solid andgas is adopted to enhance the gas and solid conversion. Thecounter-current flow pattern minimizes the back-mixing of both solid andgas. Moreover, it maintains the solid outlet of the reactor at a morereductive environment while the gas outlet of the reactor is maintainedat a more oxidative environment. As a result, the gas and solidconversions are both enhanced.

Referring back to the oxidation reaction in the second reactor in FIG.3, i.e. the oxidizer, the oxidizer converts the iron containing oxygencarrier particles from the reducer to a higher oxidation state using theexhaust gas stream of the fuel cell anode, which is rich in CO₂ and/orsteam. The presence of unconverted fuel in this stream will notparticipate in the reaction. The possible reactions include:

Fe+CO₂/H₂O=FeO+CO/H₂

3FeO+CO₂/H₂O=Fe₃O₄+CO/H₂

In certain embodiments, only a portion of the reduced oxygen carrierfrom the reducer is introduced to the oxidizer with the rest bypassingthe oxidizer and is directly sent to the combustor. By doing this, moreheat is generated from the redox block to compensate for the reactionheat required in the reducer. Alternatively, a sub-stoichiometric amountof fuel cell anode exhaust gas is sent to the oxidizer so that more heatis produced in the combustor that follows.

Although unconverted fuel may be present in the fuel cell anode exhauststream, the fuel content in this gas stream is significantly enrichedresulting from the reaction between iron/iron oxide and H₂O/CO₂.

The preferred designs of the oxidizer also include a moving bed reactorand other reactor designs that provide a countercurrent gas-solidcontacting pattern. A countercurrent flow pattern is preferred so that ahigh steam to hydrogen and CO₂ to CO conversion are achieved.

Referring back to the oxidation reaction in the third reactor in FIG. 3,i.e. the combustor, oxygen containing gas such as air and/or partiallyconverted air from the fuel cell cathode side is used to, at leastpartially, oxidize the iron in oxygen carrier generated from theoxidizer to Fe₂O₃. The reactions in the combustor include:

4FeO+O₂=2Fe₂O₃

4Fe₃O₄+O₂=6Fe₂O₃

The preferred reactor designs for the combustor include a fast fluidizedbed reactor, an entrained bed reactor, a transport bed reactor, or amechanical conveying system. The functions of the combustor include:oxidation of the oxygen carrier to a higher oxidation state; andre-circulation of the oxygen carrier to the inlet of the reducer foranother redox cycle.

FIG. 4 illustrates a schematic flow diagram of one embodiment of thereforming/water splitting block that converts gaseous fuel. In thisembodiment, a gasification system is used to convert solid fuel such ascoal, biomass, pet coke, and wax into a gaseous fuel. Sulfur in thegaseous fuel is removed using a high temperature sorbent such as thosecontaining calcium oxide, zinc oxide etc. The required sulfur level inthe gaseous fuel is <500 ppm. In preferred schemes, the sulfur level inthe gaseous fuel is reduced to <20 ppm.

The fuel gas is then introduced to the reducer in FIG. 4 as the fuel forthe redox cycles. Alternative to the gaseous fuel from the gasifier,fuels from the reformer or pyrolyzer can also be used in the redoxsystem. Gaseous fuels such as methane and light hydrocarbon can also bedirectly introduced to the redox system as the fuel.

One difference between the process and system described in Fan, PCT

Application No. WO 2010037011 and embodiments of the present inventionis that the gaseous feed for the second reactor, the oxidizer, containsfuel gas such as H₂ and CO in addition to H₂O and CO₂. In certainembodiments, the oxygen containing gas for the combustor comprises atleast a portion of the exhaust gas from the cathode.

The combustor is highly exothermic. The heat generated in the combustorcan be used to compensate for the heat required in the reducer. Thisheat can also be used to preheat the feed streams and to generate powerfor parasitic energy consumptions. The high pressure gaseous streamdischarged from the system can be used to drive expanders for gascompression.

Table 1 illustrates the mass flow of the major streams in one embodimentof the process. Table 2 illustrates the energy balance of one embodimentof the system. In this case, methane is used as the fuel. H₂O/H₂ is usedas the working fluid. The fuel cell block, which utilizes an SOFCsystem, converts the fuel (H₂) rich gas stream into 70% steam balancedwith H₂. The HHV efficiency of the process, defined as the energy in theelectricity product divided by the higher heating value of the methanefeed, is greater than 60%. In this case, substantially all of the CO₂produced is compressed to 1.52 x 10⁷ Pa (2200 psi) and is ready forsequestration.

TABLE 1 Mass Balance of the Integrated Reforming - Fuel Cell for PowerGeneration using Methane as the Fuel CO₂ H₂ rich H₂O rich Methane fromReducer stream from stream from fuel (feed, kmol/s) (kmol/s)* oxidizer(kmol/s)⁺ cell anode (kmol/s) 1.12 1.12 6.99 6.99 *the CO₂ streamcontains less than 0.5% impurities such as unconverted fuel ⁺exhaustfrom the oxidizer contains 70% H₂ and 30% steam

TABLE 2 Energy Balance of the Integrated Reforming - Fuel Cell for PowerGeneration using Methane as the Fuel Parasitic Power Power ProductionNet Methane (MW_(th)) (MWe) (MWe) Power (M)We 1000 80 700 620

In the case where coal and a coal gasifier are used, the processefficiency varies between 38 and 60% (HHV, with CO₂ capture) dependingon the type of coal and coal gasifier. When biomass is gasified and usedfor the redox system, the efficiency is 1-10% less than its coalcounterpart. Because all of the CO₂ in the biomass is captured, the netCO₂ emission from the system is negative from the life cycle analysisstandpoint.

Referring to the embodiment illustrated in FIG. 5, solid fuel such ascoal, biomass, wax, heavy residue, pet coke, and tar sands are directlyconverted in the redox system without the need for agasifier/pyrolyzer/reformer. This embodiment depicts a direct coal redoxsystem integrated with solid oxide fuel cell (SOFC) as exemplifiedherein.

Due to the high operating temperatures in a SOFC system, between about800° C. to 1000° C., a significant amount of heat is released and needsto be recovered to enhance the process efficiency. Current processdesigns usually combine SOFC and a gas turbine—steam turbine system forfull conversion of fuel to electricity. About 60%˜90% of the fuel isconverted in the SOFC first, and the remainder will be fully convertedin a gas turbine system together with a bottoming Rankine cycle.However, the system is costly because all three components, i.e., thehydrogen production system, fuel cell, and turbine system, are capitalintensive. Conventional IGCC-SOFC routes for electricity generation canreach an efficiency of at most 55%.

The direct chemical looping (DCL) process, described in Fan, PCTApplication No. WO 2010037011, converts solid fuels into hydrogen.Within the DCL system, an iron oxide based oxygen carrier circulatesamong three reactors which are the reducer, the oxidizer and thecombustor. In the reducer, coal and/or biomass is gasified to CO₂ andH₂O by Fe₂O₃ containing particles which are reduced to Fe and FeO. Aportion of the reduced particles react with steam in the oxidizer toproduce hydrogen, while the remaining reduced particles together withthe partially oxidized particles from the oxidizer, are fed to thecombustor. Finally, Fe₂O₃ containing particles are regenerated andrecycled back by combusting with oxygen containing gases such aspressurized air. The heat, released in the combustor and carried over tothe reducer by the iron oxides, can fully compensate for any heatdeficit in the system. By the DCL system, hydrogen and carbon dioxideare generated in different reactors, which saves a considerable amountof energy by eliminating the need for product separation. Also, it savesequipment investment costs on CO₂ removal and air separation units. TheDCL system can produce hydrogen at an efficiency of 70-85% from coal and60-75% from biomass.

In this embodiment, we integrate the DCL system and SOFC system for highefficiency electricity generation from coal. The DCL-SOFC process andsystem have a number of configurations, either at high pressure or lowpressure. Specifically, we describe the embodiment where the oxidizerand anode are integrated within a closed loop of hydrogen and steam asshown in FIGS. 5 and 9.

1000 MW thermal input is considered, and accordingly 131.8 tonne/hr ofbituminous coal is processed in the DCL-SOFC system. Coal is firstpulverized into proper size particles and then dried to 5% moisture from7.23% by the flue gas. In the

DCL system, a moving bed design is adopted for both the reducer and theoxidizer. About 3549.5 tonne/hr oxygen carrier, containing 45.6% Fe₂O₃and 54.4% Al₂O₃ (as inert) by weight, is fed into the top of thereducer, and the coal is injected from the middle part of the reducer.In the moving bed reducer, solid flows downward while gas ascendsupward. The countercurrent design can fully convert coal into CO₂ andH₂O at 900° C., 1.01×10⁵ Pa (1 atm). Iron oxide is reduced to the formof Fe, FeO and a trace of FeS. 71.5% of the reduced iron particles areused for hydrogen generation in the oxidizer, and the other 28.5% arecombusted in the combustor. The oxidizer operates at 850° C., convertinga gaseous mixture of 90.4% H₂O and 9.6% H₂ by mole into a mixture of35.9% H₂O and 64.1% H₂ and ppm level of H₂S. The gaseous mixture is thenfed to the anode of a sulfur tolerant SOFC for electricity generation.At the same time, Fe and FeO will be oxidized to Fe₃O₄, which flows tothe combustor for Fe₂O₃ regeneration.

An air blower drives 1992 tonne/hr of air to feed the DCL-SOFC system.The air is preheated up to 900° C. in the HRSG section, and then goes tothe cathode of the SOFC device. 30% of the oxygen and 85% of thehydrogen are consumed in SOFC operating at 900° C. The spent air is usedin the combustor to regenerate Fe₂O₃ at 1280 ° C. The spenthydrogen/steam mixture will then be cooled to about 240° C. forsubsequent sulfur removal unit. Only a small amount of steam will bemade up to the hydrogen/steam mixture before it recycles back to theoxidizer.

During the DCL-SOFC process, >99% pure CO₂ is obtained by simplecondensation followed by compression to 1.37×10⁷ Pa (>135 atm) forgreenhouse gas control. The compression step consumes about 35.8 MW ofwork. The other pollutants such as Cl, S, and Hg can either beco-sequestered with CO₂ or removed by conventional techniques. Ash canbe removed from the oxygen carrier by a cyclone positioned before thereducer.

Table 3 summarizes the flow of the main process streams. As a result ofthe integration of the DCL and SOFC, 535 MW of electricity can beproduced by the DCL-SOFC system, and 96 MW of electricity can begenerated from the steam turbine system by recovering low grade heat.The overall process can produce electricity of 640 MW with CO₂compression, this is equal to a coal to power efficiency of 64% (HHV).The illustrated example can be further optimized to achieve greater than70% efficiency.

The DCL-SOFC system can convert a wide range of combinations of coal andbiomass to electricity with high efficiency. Possible designs alsoincludes low pressure and temperature operation for the working fluid(the mixture of hydrogen and steam). H₂S in the hydrogen/steam mixturecan be also removed before the SOFC with hot gas clean up unit. It isnoted that when feeding the system with low sulfur fuel (approximatelyless than 0.2% by weight) such as biomass, no sulfur removal unit isneeded.

TABLE 3 Process Flowsheet for the DCL-SOFC Process Stream 0 1 2 3 4 5 67 8 9 Temperature 30 1280 901 901 850 1279.6 901 30 159.9 1279.6 ° C.Pressure atm 30 30 1 1 30 16 1 1 135 16 Mass Flow 131.878 3549.4592336.231 931.225 2475.293 3549.459 402.831 348.195 348.195 1709.881tonne/hr Volume Flow 92.642 871.979 855.716 341.089 667.534 871.9791.05E+06 195768.7 1754.619 481020.7 cum/hr Density 88.868 254.118170.438 170.438 231.49 254.118 0.024 0.111 12.388 0.222 lb/cuft MassFlow tonne/hr H₂O 0 0 0 0 0 0 53.058 0 0 0 CO₂ 0 0 0 0 0 0 348.195348.195 348.195 0 O₂ 0 0 0 0 0 0 0 0 0 180.486 N₂ 0 0 0 0 0 0 1.19 0 01526.995 H₂S 0 0 0 0 0 0 0 0 0 0 H₂ 0 0 0 0 0 0 0 0 0 0 SO₂ 0 0 0 0 0 00.388 0 0 0.141 SO₃ 0 0 0 0 0 0 0 0 0 0.006 NO 0 0 0 0 0 0 0 0 0 2.212NO₂ 0 0 0 0 0 0 0 0 0 0.039 Fe 0 0 326.047 129.963 0 0 0 0 0 0Fe_(0.947)O 0 0 630.558 251.341 391.333 0 0 0 0 0 Fe₃O₄ 0 0 0 0 704.7990 0 0 0 0 Fe₂O₃ 0 1620.562 0 0 0 1620.562 0 0 0 0 Fe_(0.877)S 0 0 0.4640.185 0 0 0 0 0 0 Al₂O₃ 0 1928.897 1379.162 549.736 1379.162 1928.897 00 0 0 COAL 131.878 0 0 0 0 0 0 0 0 0 Stream 10 11 12 13 14 15 16 17 1819 Temperature 120 25 59.1 900 900 850 900 240 30 600 ° C. Pressure 16 12 2 2 30 30 30 124 124 atm Mass Flow 1709.881 1992.014 1992.014 1992.0141852.821 123.993 263.185 263.185 360.398 360.398 tonne/hr Volume Flow121911.2 1.69E+06 940966.6 3.32E+06 3.12E+06 49218.57 51174.16 20307.05425.414 10786.28 cum/hr Density 0.876 0.074 0.132 0.037 0.037 0.1570.321 0.809 52.887 2.086 lb/cuft Mass Flow tonne/hr H₂O 0 0 0 0 0103.165 259.895 259.895 360.398 360.398 CO₂ 0 0 0 0 0 0 0 0 0 0 O₂180.486 463.974 463.974 463.974 324.782 0 0 0 0 0 N₂ 1526.995 1528.041528.04 1528.04 1528.04 0 0 0 0 0 H₂S 0 0 0 0 0 0.195 0.195 0.195 0 0 H₂0 0 0 0 0 20.633 3.095 3.095 0 0 SO₂ 0.141 0 0 0 0 0 0 0 0 0 SO₃ 0.006 00 0 0 0 0 0 0 0 NO 2.212 0 0 0 0 0 0 0 0 0 NO₂ 0.039 0 0 0 0 0 0 0 0 0

TABLE 4 Coal to Electricity Process Configurations and ProcessEfficiencies Process Configuration DCL-SOFC DCL-SOFC with closedDCL-SOFC with oxidizer-anode with further Conventional DCL-SOFTcombustor- loop and heat Gasification - without cathode combustorintegration of WGS-SOFC integration integration integration reducerEfficiency 38-48% 50-55% 51-57% 58-64% 66-71% (% HHV with CO₂ capture)

Although the DCL-SOFC system and process exemplified in this embodimentis specific to working fluid compositions, type of reforming/watersplitting block, and fuel cell block, the choices of aforementionedparameters have a large degree of freedom. For instance, CO and CO₂ canbe used instead of H₂/H₂O as the working fluid. The variousconfigurations described in Fan, PCT Application No. WO 2010037011 canbe used in the reforming/water splitting block. Other fuel cells such asmolten carbonate fuel cell (MCFC) can also be integrated with the DCLsystem. In this case, a portion of the CO₂ generated from the DCLreducer is injected to the cathode side of the MCFC to facilitate theconversion. In addition, the DCL system can be configured so that thegaseous exhaust from the reducer is not fully converted. In this case,the unconverted fuel is sent to another fuel cell and/or oxygenpolishing step prior to obtaining a concentrated CO₂ stream (see FIG.10). When all the reduced oxygen carrier particles are used for hydrogenproduction, i.e. the split ratio for direct combustion is 0, andassuming high grade heat in the fluegas from the combustor can be usedto heat up the reducer, the electricity generation efficiency can reach70% with CO₂ compression. Table 4 shows the several configurations andcorresponding power generation efficiencies.

FIG. 6 illustrates an embodiment in which a calcium sorbent enhances thereforming process and is used as the reforming/water splitting block. Inthis case, the fuel is reformed/shifted to H₂ with the presence ofCaO/Ca(OH)₂ sorbent and steam/steam rich exhaust gas from the fuel cellanode:

CaO+C_(x)H_(y)+H₂O→CaCO₃+H₂

The spent sorbent is then regenerated at high temperatures using thewaste heat from the system in the calciner:

CaCO₃=CaO+CO₂

A hydration step is optionally added to reactivate the sorbent. Theconcentrated CO₂ from the calciner is then compressed and sequestered.In this case, a portion of the working fluid can be split to avoidaccumulation of the working fluid.

FIG. 7 illustrates the option of using a membrane enhancedreformer/water gas shift reactor as the reforming/water splitting block.In this embodiment, the fuel is reformed/shifted in the reformer, andCO₂ is simultaneously removed from the membrane. The retentate side ofreformer enriches the working fluid with reformed fuel, while thepermeate side produces concentrated CO₂.

FIG. 8 illustrates an embodiment showing the integration of a zinc oxidewater splitting cycle and the fuel cell. In this embodiment, thermalenergy from a solar or nuclear source is used to facilitate the zincoxide based water splitting cycle. The hydrogen obtained from thesplitting of water is used to enrich the working fluid comprising H₂Oand H₂.

It will be apparent to those skilled in the art that various changes maybe made without departing from the scope of the invention which is notconsidered limited to the specific embodiments described in thespecification and drawings, but is only limited by the scope of theappended claims.

1. A system for converting carbonaceous fuel or thermal energy intoelectricity comprising: a reforming/water splitting block for convertinga steam and/or CO₂ rich gas stream and carbonaceous fuel and/or thermalenergy into a fuel (H₂ and/or CO) rich gas stream and an exhaust gasstream; a fuel cell block for converting the fuel rich gas stream and anoxygen containing gas stream into a lean fuel gas stream and a spentoxygen containing gas stream from anode and cathode, respectively; and aclosed loop between the reforming/water splitting block and fuel cellblock. 3-22. (canceled)
 23. A system as claimed in claim 1 in which thefuel cell block comprises a solid oxide fuel cell or a molten carbonatefuel cell.
 24. A system as claimed in claim 1 in which the carbonaceousfuel comprises syngas, carbon monoxide, methane rich gas, lighthydrocarbons, liquid carbonaceous fuels, coal, biomass, tar sand, oilshale, petroleum coke, heavy liquid hydrocarbons, wax, and mixturesthereof.
 25. A system as claimed in claim 1 in which less than 10% ofthe fuel rich or steam/CO₂ rich gas stream is purged.
 26. A system asclaimed in claim 1 in which the thermal energy comprises either solar ornuclear energy.
 27. A system as claimed in claim 1 in which the fuelrich stream is produced through water splitting, electrolysis, sorbentor membrane enhanced reforming and/or water gas shift reaction, orsteam-iron reaction.
 28. A system as claimed in claim 1 in which theredox property of a metal oxide particle is used to assist in thecarbonaceous fuel conversion.
 29. A system as claimed in claim 1 inwhich a calcium oxide sorbent is used to enhance the carbonaceous fuelconversion.